Conversion of residual oil to gasoline



United States Patent Office 3,072,560 Patented Jan. 8, 1963 This invention relates to a process for converting residual oils to gasoline fractions. More particularly, the invention is directed to a process wherein a heavy oil such as a reduced crude petroleum stock is first passed through a coking unit and wherein gas oil fractions recovered from the coker are denitried and then converted to gasoline fractions in yields approaching 100 volume percent as they are admixed with hydrogen and passed over a hydrocracking catalyst at relatively low temperatures.

As employed herein, the terms residual oil and residual stock are used to designate materials such as reduced crude petroleum oils recovered asrbottoms from a distillation column to which crude petroleum is fed or from a vacuum stripping unit utilizing said bottoms as feed; asphaltic residues recovered from a propane or other deasphalting unit; residuum from thermal or catalytic cracking operations; certain of the heavier gas oils; and residual stocks as derived from hydrocarbon-bearing materials of non-petroleum origin such as shale oil, gilsonite, coal, or lignite. Of such stocks, those presenting themselves in largest volume are the reduced crude petroleum oils. Accordingly, for convenience of description, the invention will hereinafter be more particulary described as it relates to the conversion of such oils, though it should be understood that other residual stocks may be processed with like results in a generally similar fashion.

In the operation of a reiinery, residual stocks of one type or another are produced in substantial volume. While portions of said stocks may be used for bunker or other fuel oil purposes, such products have a relatively low value, and it is therefore the practice to convert the residual stocks to gasoline fractions insofar as possible. However, the yield of gasoline which can be obtained from residual oils by a practice of available refinery methods is rather low. This is borne out by the data presented in Table I below as calculated from a representative run wherein a reduced crude oil is first passed through a coking unit, with the resulting gas oil fractions from the coker then being processed inthe conventional manner using available thermal and catalytic cracking facilities, a catalytic reformer and an alkylation unit wherein isobutane is reacted with C3 and C4 olelins to produce so-called alkylate. More specifically, this run assumes an initial production from the reduced crude of 10,000 b./d. (barrels per operating day) of gasoil from the coking unit, said gas oil boiling over a range from about 400 to 800 F. This gas oil is passed to a catalytic cracking unit from which are recovered various light gas and gasoline fractions, as well as both light and heavy cycle oils. To keep the operation in balance, 1000 b./d. of the light cycle oil so produced (which is highly refractory and resistant to further cracking) is diverted to fuel oil purposes, while the balance of the light cycle oil (1250 b./d.) is fed along with the heavy cycle oil (2250 b./d.) to a thermal cracking unit from which are recovered various light gas and gasoline fractions. The Cq-- gasoline portion (850 b./d.) so recovered, boiling from about 180 to 400 F., is passed to a catalytic reformer. lsobutane recovered from the various units is converted to gasoline by alkylation with an equivalent amount of Ca-C, product olelins. Summing up all of the products obtained in thus converting 10,000 b./d. of coker gas oil, the following distribution of products is obtained.

TABLE I Product: B./d. Total gasoline (F-l octane number, plus 3 ml.

TEL=96.7 5310 C1-C3 non-olefin product (equiv. liq. yield) 1160 `C3-C., oleiins (excess over alkylate requirement) Fuel oil (175 SSF at 122 F.) 3790 It will be observed from the above table that only 5310 barrels of gasoline are produced from 10,000 barrels of col-2er gas oil, although the gasoline product values approach 6500 barrels as additional isobutane is brought in from external sources and combined with the excess of oletins, as shown, to form more alkylate.

lt is an object of this invention to provide a process wherein the yield of gasoline is much higher than that shown above and approaches 100 volume percent, based on Coker gas oil. A further object is to provide a process wherein the gasoline so prepared is characterized by a significantly higher overall leaded octane rating than has heretofore been possible.

The present invention is based on the discovery that the foregoing objects can be lachieved by the practice of a novel combination of hydrocarbon conversion steps wherein the reduced crude petroleum oil or other residual stock employed is first converted to lighter boiling fractions and coke by passage through a coking unit under coking conditions, and wherein nitrogen-containing gas oil fractions recovered from the coker are first denitried and then converted in high yield to lower boiling gasoline product fractions by passage, in admixture with at least 2000 scf. of hydrogen per barrel of feed, over a hydrocracking catalyst at temperatures of from about 350 to 700 F. and pressures of at least 400 p.s.i.g., said hydrocracking catalyst comprising an acidic material having hydrogenating characteristics and high cracking activity. ln the preferred practice of the invention, gasoline fractions so produced which boil above about 165 F. (e.g., a F-400o F. cut) are passed through a catalytic reformer under reforming conditions. Portions of the product stream from the hydrocracking catalyst boiling above the gasoline range can be converted to gasoline fractions as they are recycled to the catalyst. Alternatively, said higher boiling fractions may be diverted, in whole or in part, to jet or other fuel oil end uses.

The advantages to be gained by a practice of this invention can be shown as the process is utilized in connection with the same coker gas oil feed as employed in the run described above. In this case, the coker gas oil is first reduced in total nitrogen content to a level below 10 p.p m. by passage, along with hydrogen, over a hydrolining catalyst under hydroning conditions, said conditions resulting in a generally selective decomposition of nitrogen-containing and sulfur-containing compounds present in the feed. The resulting low-nitrogen gas oil obtained from the hydrofining unit is then combined with 6500 s.c.f. H2 per barrel of feed and passed, at an LHSV of 1, a pressure of i500 p.s.i.g. and a temperature ranging from about 550 to 700 F., over a hydrocracking catalyst incorporating a hydrogenating component (eg, nickel sulfide) carried on a synthetically prepared cracking support comprising silica-alumina containing approximately of the silica component. In this operation, the gas oil is converted to gasoline fractions at a per-pass conversion of approximately 50-60%, said conversion being attended by a hydrogen consumption of approximately 1000 s.c.f. per barrel of feed converted to synthetic product, i.e., that boiling below the initial boiling point of the feed. Of the total effluent stream from the hydrocrack-h ing catalyst, those portions boiling above 400 F. are recycled back over the catalyst, while the Cq-iportions of the effluent (i.e., those boiling from about 180 to 400 F.) are passed, along with added hydrogen, over a reforming catalyst under reforming conditions, some 6300 b./d. of gasoline being reformed in this fashion, The following products are obtained from the above operation assuming a feed rate of 10,000 b./d. of coker gas oil.

TABLE II Product: B d.

Total gasoline (F-l octane rating, plus 3 ml.

TEL=100-0 9500 Isobutane (available as feed to alkylation plant) 1400 C1-C3 products (equivalent liquid yield) 1150 Fuel oil 0 It will be seen from Table II that extremely high yields of gasoline are obtained when the present invention is so practiced as to hydrocrack all the gas oil produced by the coking unit, the gasoline yield being even better (increased about 2500 b./d.) when C3-C4 olefins are brought in from another refinery source to alkylate the isobutane which is formed during the hydrocracking step. It is also apparent that the gasoline produced has an extremely goed octane number.

The present invention also lends itself well to the method wherein, of the total gas oil recovered from the coker, only the lighter fractions, i.e., those boiling within a range of from about 325 to 600 F., are denitried and subsequently hydrocracked. Thus, in the operation referred to above wherein the coking unit has a production of 10,000 b./d. of gas oil, approximately 3500 b./d. thereof are of the light variety. Denitrification and subsequent conversion of this light gas oil in the hydrocracker unit under the conditions shown above gives the following product distribution.

TABLE III Product: B./d.

Total gasoline (F-l octane rating, plus 3 ml.

TEL=100.2 3534 -Isobutane (available as feed to alkylation plant) 385 C1-C3 products (equivalent liq. yield) 100 In the event that the invention is so practiced as to feed only the lighter coker gas oil portions to the hydrocracker, the heavier gas oil portions will normally be fed to a cracking unit of the thermal or catalytic type. When this procedure is followed, it has been found that especially good results are obtained when the effluent from the cracking unit is so worked up as to recover a light cycle oil fraction, i.e., one boiling within a range from about 350 to 650 F., with the portion so recovered being denitried and then passed over the hydrocracking catalyst, preferably in conjunction with the dentriied light gas oil from the coker. When operating in this fashion, the heavier cycle oil portions ofthe effluent from the cracking unit can either be recycled to extinction therein, or they may be processed by passage through other appropriate refinery units.

Table IV below presents data showing the nature of the products obtained when the denitrified light coker gas oil (3500 b./ d.) is hydrocracked and the heavy coker gas oil (6500 b./d is converted in a catalytic cracking unit to light catalytic cycle oil (2000 b./d.), gasoline and lighter products, the light cycle oil recovered from the catalytic cracker being denitrifed and then passed to the hydrocracking unit along with the denitrilied light coker gas oil. In this operation, some 4000 b./d. of heavy gasoline from the hydrocracker are passed over the reforming catalyst.

i TABLE IV Product: B./d.

Total gasoline 1 (F-l octane rating plus 3 ml.

TEL=99.5) 9500 `C-C.,t olens (excess over alkylate requirement) 300 Cl-CS (equivalent liq. yield) 1000 lIncludes 1000 b./r1 produced in alkylation plant from i-Ci and Ca-Ci olelins produced during the process.

As set forth above, the first step of the present conversion process involves passing the residual feed through a coking unit under conventional coking conditions. The units presently available for this purpose are of either the delayed or fluid variety. When using a delayed Coker, the reduced crude oil or other residual feed is heated to about 750 950 F. and then fed to one of two or more vertical, insulated coke drums. The drums are connected by valves so that they may be put on stream for filling, and then taken off stream for coke removal as the amount of coke formed therein builds up to maximum capacity. The temperature in the drum will ordinarily be of the order of 775S50 F. and the pressure 40-60 p.s.i.g. Hot vapors from the coke drum pass to a fractionator where gas, various gasoline fractions, light gas oil, and heavy gas oil are separated, a portion of the heavy gas oil being recycled to the furnace inlet, if desired.

When resort is had to the use of a fluid type of coking unit, a residual feed is sprayed into a chamber for contact with hot particulate solids maintained in a so-called liuidized condition. Upon contact with the solids, the oil undergoes pyrolysis, evolving lighter hydrocarbons and depositing carbonaceous residue on the solid particles, causing them to grow in size. The necessary heat for the pyrolysis is supplied by circulating a stream of the uidized solids through an external heating or combustion zone and then passing the resulting extremely hot coke particles back to the fluidized coking Zone proper for contact with incoming feed. The necessary fluidization of the various particulate streams in the unit can be obtained by the use of steam, although it is also possible to use various light hydrocarbon gases for this purpose. The vaporous products formed in the fluidized coking chamber are withdrawn through a cyclone type of separation unit for removal of entrained particulate solids, with the product stream then being worked up in the conventional fashion to recover various normally gaseous and gasoline streams, as well as both light and heavy gas oils. Portions of the latter may be recycled to the unit, if desired.

Depending on the nature of the particular coking unit employed, the temperature maintained therein and such factors as the cut point selected for gasoline fractions and the extent to which heavy gas oil fractions may be recycled to the coking zone, coker gas oil fractions boiling over a range from about 400 F. to 1000 F., or even somewhat higher, may be obtained. Of such gas oils, those boiling within a range from about 350 to 650 F. are conventionally designated as light gas oils, while those of higher boiling range are termed heavy gas oils. The present invention can be employed successfully by hydrocracking the entire gas oil spectrum from the coker or any portion thereof. However, in the preferred practice of the invention, of the gas oil fractions recovered from the coker only those of the light variety are denitried and sent to the hydrocracker; the heavier gas fractions are subjected to treatment in a conventional thermal cracking unit or in a catalytic cracker of the moving bed or fluid variety.

As noted above, the hydrocarbon feed streams passed to the hydrocracking zone should be subjected to a preliminary denitriiication treatment, the latter being practiced with such portions of the coker gas oil as are to be hydrocracked, as well as with light cycle oil stocks and, if necessary, with any other feed stocks which may be passed over the hydrocracking catalyst.

lt has been found that nitrogen-containingcompounds present in the feed passed over the hydrocracking catalyst rapidly reduce the activity thereof. While catalyst activity can be temporarily maintained at a relatively high level in the presence'of nitrogen compounds by resort to high temperatures, i.e., those falling above 700 F., this expedient defeats the purposes of the present invention since it leads to rapid catalyst coking unless extreme pressures are employed, and thus has the effect of shortening the onstream portion of a given run to a period of uneconomic duration. High temperatures are also undesirable, since they induce over-cracking and thus greatly increase the amount of light gases produced during the hydrocracking step. In keeping with the process of this invention, it has been found that relatively low temperatures may be employed over the hydrocracking catalyst when the total nitrogen content of the feed thereto is maintained below 1,0 p.p.m., with even further benefits being obtained when the denitrication treatment practiced is such as to reduce the total nitrogen content of the feed to a level of from to 2 p.p.m. g

Any available denitrification method can be 'employed which is effective to reduce the nitrogen content of the par-y ticular feed stocks to the desired level. For example, some feeds can be denitrified by intimately contacting the samev with various acidic media such as liquid acids (H2804 or the like) or with various solid acidic materials which are capable of selectively absorbing nitrogen-containing compounds present in the feed. However, a preferred denitrification method of broader utility involves passing the feed, along with at least 500 s.c.f. of hydrogen per barrel thereof, over a sulfur-resistant hydrogenation catalyst at temperatures of from about 450 to 800 F., pressures of at least 300 p.s.i.g., and liquid hourly space velocities (LHSV) of from about 0.3 to 5, the conditions being so chosen that little cracking of the feed takes place other than that of the nitrogenand sulfur-containing cornpounds present. While any of the known sulfur-resistant hydrogenation catalysts can be used, the preferred catalysts of this category have as their main active ingredient one or more of the transition metals such as cobalt, molybdenu-m, nickel, or tungsten, or oxides or suldes of such metals. These materials may be used in a variety of combinations with or without the use of various known stabilizers and promoters. Moreover, these catalysts may be employed either alone or in combination with various conventional supporting materials such as charcoal, fullers earth, kieselguhr, silica gel, alumina, bauxite, or magnesia. A representative effective denitriication catalyst for use in the present invention is one embodying an alumina support and containing molybdenum and/or tunsten in the sulfide or oxide form, in an amount of about to 60% expressed as Mo or W, together with oxides or sultides of cobalt and/or nickel, the latter materials being present in the amount of from about 1 to 20%, expressed as Ni or Co. This method of denitrification can be referred to as a hydrofining treatment.

The effluent obtained from the denitriiication treatment is treated, in accordance with the methods presently known in the art, so as to remove ammonia and at least part of hydrogen sulfide which may be present. The denitried hydrocarbon feed stock is then ready to be hydrocracked by passage, along with added hydrogen, over a hydrocracking catalyst at elevated pressures and at temperatures ranging from about 350 to 700 F.

The catalyst employed in the hydrocrackng unit is an acidic material having hydrogenating characteristics and high cracking activity. It is made up of a hydrogenating component together with a material having a high degree of cracking activity either per se or when combined with the material employed to provide a hydrogenating component of the catalyst. In this connection, the term high cracking activity is employed herein to designate those catalysts having activity equivalent toa Cat. A value 6 of at least or a quinoline number of at least 20 (Journal Am. Chem. Society, 72, 1554 (1950)). In the case of catalysts not adapted to withstand the conditions employed in such tests, generally comparable, minimal cracking activity values can be determined by other methods known in the art.

Broadly speaking, the hydrogenating component of the catalyst may comprise a compound of one or more of the metals in groups I(B), II(B), V, VI, VII and VIII of the periodic table, said compound being one which is not readily reduced to the corresponding metal form under the reducing conditions prevailing in the hydrocracking zone. Representative compounds of this character include oxides and suldes of molybdenum, tungsten, chromium, rhenium and zinc, as well as sullides of cobalt, nickel, copper and cadmium. Other suitable hydrogenating components coming Within the category of non-readily reducible compounds are complexes `of the various metals of the defined groups such, for example, as cobalt-chromium and nickel-chromium, representative preparations of this character being described in U.S. Patent No. 2,899,287. If desired, more than one hydrogenating component may be present. The amount of the hydrogenating component may be varied within relatively Wide limits of from about 0.1 to 35% or more, based on the Weight of the entire catalyst composition.

The cracking component of the hydrocracking catalyst may be selected from a variety of solid or liquid materials of the type having good cracking activity. Among solid compositions which can be used are the various siliceous cracking catalysts; those wherein alumina is chemically bonded to aluminum chloride; tluorided magnesium oxide; and aluminum chloride, particularly when contained within the pores of a support such as charcoal so as to reduce vaporization of the AlCla. Representative liquid catalysts having a high degree of cracking activity are hydrogen fluoride-boron triuoride compositions, titanium trichloride, and aluminum chloride as contained in a suitable hydrocarbon vehicle along with HC1. Y

In general, it is preferred to employ a solid siliceous material as the cracking component of the catalyst. For example, there may be used composites of silica-alumina, silica-magnesium, silica-alumina-zirconia, acid treated clays and the like, as well as synthetic metal aluminum silicates (including synthetic chabazites normally referred to as molecular sieves) which have been found to impart the necessary degree of cracking activity to the cata lyst. Particularly preferred siliceous catalyst components are synthetically prepared silica-alumina compositions having a silica content in the range of from about 40 to 99% by weight.

Particularly good results from the standpoint of high per-pass conversion, even at relatively low operating ternperatures, coupled with a high ratio of iso to normal parafns in the synthetic product stream from the hydrocracking catalyst, are obtained With catalysts comprising a total of from about 0.1 to 35 wt. percent of at least one compound selected from the group consisting of cobalt sulfide and nickel sulfide, said compounds being deposited on the aforementioned synthetically prepared silicaalumina composites. Of these catalysts, those containing nickel sulfide are found to have the highest activity. Catalysts in this group can be readily regenerated, if desired, by conventional burning techniques.

The following hydrocracking catalysts are representa- Ytive of those which are adapted to be used in a practice of the present invention, the support in each case being a synthetically prepared silica-alumina composite containing about 87-90% silica and having a Cat. A value of approximately 46.

Nickel Sulfde (6% Ni) 0n Silica-Alumina This catalyst (No. A) was prepared by impregnating the silica-alumina support, present in the form of small beads, with a solution of nickel nitrate in a concentration suicient to provide the catalyst lwith 6 wt. percent nickel on a dry basis. The catalyst was dried at 600 F. and was then thermactivated by contact for 2.2 hours with a Stream of hot air at an average temperature of 1427 F., said thermactivation treatment fo-rming the subject of application Serial No. 794,109, tiled February 18, 1959, and now abandoned. The catalyst was then cooled and reduced by contact with a stream of hydrogen, rst at atmospheric pressures as the catalyst was heated from 60 to 570 F. at a rate of 100 F. per hour, and thereafter at 1500 p.s.i.g. and 570 F. for one hour. The metallic nickel present on the catalyst was then converted to the sulfide form by contacting the catalyst with a solution of iso-propyl mercaptan (10 wt. percent) in hexane, hydrogen being present in an amount such as to give the equivalent of 2 wt. percent H25 in the gas stream passed Awer the catalyst. This sulliding treatment was continued for 31/2 hours at 1500 p.s.i.g. and 570 F., a treatment which provided the catalyst with a 2.6-fold excess of sulfur over the amount theoretically required to convert all of the nickel to nickel sulfide.

NckelSulde (3.6% Ni) on Silica-Alumina This catalyst (No. 425-2) was prepared by impregnating l1 liters of a crushed SiO2-Al203 aggregate with 2896.9 grams of Ni(NO3)2-6H2O, dissolved in enough water to make 880 milliliters total solution, following which the material was held for 24 hours at 70 F. The catalyst was then dried for l hours at 250 F. and thereafter calcined at 1000 F. for 10 hours. The calcined material was reduced in an atmosphere of hydrogen at 580 F. and 1200 p.s.i.g., following which the resulting nickel-bearing catalyst was sulfided in an atmosphere containing 8% H25 in hydrogen at 1200 p.s.i.g. and 580 F., thereby converting essentially all the nickel to nickel sulfide.

Nickel Sulfde (2.5% Ni) 0n Silica-Alumina This catalyst (No. 316) was prepared by impregnating 11 liters of a crushed SiO2-Al203 aggregate with a solution prepared by mixing 1500 milliliters water and 500 milliliters of ammonium hydroxde solution with 1082 grams of ethylenediamine tetracetic acid (EDTA) and 469 grams of nickel carbonate, the solution being made up to a total of 4000 milliliters with water. The impregnated material was held for a period of 24 hours at 70 F., following which it was centrifuged and calcined for l0 hours at 1000 F. in air to convert the nickel chelate to nickel oxide. The catalyst was then reduced in an atmosphere of hydrogen at 650 F. and 1200 p.s.i.g. and suliided in situ in the reactor by the use of a feed stream made up of a catalytic cycle oil (49 volume percent aromatics) to which 0.1% by volume of dimethyl disulfide had been added at a pressure of 1200 p.s.i.g., and in the presence of approximately 6500 s.c.f. H2 per barrel of feed.

Cobalt Sulfde (4% Co) on Silica-Alumina This catalyst (No. 248-2) was prepared by impregnating 2000 milliliters of a crushed SiO2-A12O3 aggregate with 1500 milliliters of an aqueous solution containing 172.5 milliliters ammonium hydroxide solution and 373 grams EDTA along with 168 grams cobalt carbonate, the solution being heated until bubbling ceased before being added to the silica-alumina material which, in turn, had previously been dried for 24 hours at 400 F. Following impregnation, the catalyst was centrifuged and calcined for four hours at 1000 F., thus yielding a material having an amount of cobalt oxide equivalent to 2.2% weight percent Co. A second impregnating solution was then made up as above, using 150.2 grams cobalt carbonate, 334 grams EDTA and 154 milliliters of ammonium hydroxide and added to the catalyst. Following a holding period of 24 hours at 70 F., the catalyst was centrifuged and calcined for hours at 1000 F.

The calcined product so obtained was then alternately reduced in hydrogen and oxided in air (repeating the cycle 5 times) at 1000 F. and 1200 p.s.i.g. The catalyst was then sulded by treatment with an excess of a mixture comprising 10% by volume of dimethyl disulfide in mixed hexanes at 1200 p.s.i.g. and 675 F., hydrogen also being present in the amount of about 6500 s.c.f. per barrel of feed.

Cobalt Sulfde (2% Co) and Chromium Sulfde (3.53% Cr) on Silica-Alumina This catalyst (No. 174-5) was prepared by forming an aqueous slurry with 1130 grams of the chelate of chromium and EDTA, to which slurry was added 196 grams of cobalt carbonate, the solution being then stirred until bubbling action ceased and made up to 1779 milliliters. This solution was warmed to 140 F. and added to 2280 milliliters of the crushed SiO2-Al203 aggregate. The resulting material was then held for 24 hours at 140 F., following which it was centrifuged and calcined 10 hours at 1000 F. The calcined product was reduced in an atmosphere of hydrogen at 1200 p.s.i.g. and 675 F., following which the cobalt and chromium metals present were converted to sulides by treatment with an excess of a solution comprising 10% by volume of dimethyl disulfide in mixed hexanes at 1200 p.s.i.g. and 675 F., hydrogen also being present in the amount of 6500 s.c.f. per barrel of feed.

Molybdenum Sulfde (2% Mo) on Silica-Alumina This catalyst (No. 226) was prepared by forming 530 milliliters of an ammoniacal solution containing 41.4 grams of ammonium molybdate. This solution was then added to the crushed SiO2-A12O3 aggregate, previously dried for 24 hours at 400 F., in an amount sufficient to yield a dried product containing the equivalent of 2 weight percent Mo. After being held for 24 hours at 70 F., the impregnated material was centrifuged and calcined for 5 hours at 1000 F. It was then reduced in an atmosphere of hydrogen at 1200 p.s.i.g. and 650 F., following which it was sulfided in situ by treatment under these same conditions of temperature and hydrogen pressure with a hydrotined cycle oil (49% aromatics) containing 1% by volume dimethyl disulfide.

Nickel Sulfde (1% Ni) and Molybdenum Sulde (1% Mo) on Silica-Alumina This catalyst (No. 296) was prepared in the following manner. 28.6 milliliters of ammonia were mixed with milliliters water and added to 49.3 grams EDTA, and to this solution was added 22.3 grams of nickel carbonate. After being heated to evolve carbon dioxide, this solution was mixed with another solution prepared by dissolving 78.7 grams of ammonium molybdate in a mixture of 80 milliliters of ammonia hydroxide and 80 milliliters of water. The resulting solution, on being made up to 480 milliliters by the addition of water, was then used to impregnate 600 milliliters of the crushed SiO2-Al203 aggregate. The impregated material, after being held for 24 hours at 70 F., was centrifuged and calcined for a period of 10 hours at 1000 F. It was then reduced in an atmosphere of hydrogen at 1200 p.s.i.g. and 650 F., following which it was sultided under these same conditions of temperature and hydrogen pressure with a solution containing 10 volume percent dimethyl disuliide in mixed hexanes.

Returning now to a general teaching of the present invention, hydrocracking is effected by passing the denitried feed stock, in admixture with at least 500 s.c.f.

of hydrogen per barrel of total feed (including both fresh' as Well as recycle feed) over the hydrocracking catalyst at temperatures of from about 350 to 700 F. and at pressures of at least 400 p.s.i.g. At least 250 s.c.f. and normally from about 500 to 1500 s.c.f. of hydrogen are consumed in the hydrocracking reaction zone per barrel 9 f total feed converted to synthetic products, i.e., those boiling below the initial boiling point of the feed to this zone.

As indicated above, the pressures employed in the hydrocracking zone are in excess of 400 p.s.i.g., and they may range upwardly to as high as 3000 p.s.i.g. or more, wlth a preferred range being from about 500 to 2000 p.s.1.g. Y

Generally, the feed may be introduced into the hydrocracking zone at a liquid hourly space velocity (LHSV) of from about 0.2 to volumes of hydrocarbon (calculated as liquid) per superlicial volume of catalyst, with a preferred rate being from about 0.5 to 3 LHSV.

One of the most advantageous aspects of the subject process is that the average reaction temperature over the hydrocrackingcatalyst can be maintained bel-ow about 700 F. while still obtaining high per-pass conversions overruns extending for several months, this without catalyst regeneration. The importanceof such low temperatureoperations is reflected in the production of extremely low yields of C1-C3 light gases and inthe formation of a synthetic product having iso to normal parain ratios far in excess of thermodynamic equilibrium values, said ratios being higher than those observed when operating at temperatures above 700 F. In the preferred practice of this invention, the temperature at which the hydrocracking reaction is initiated when placing a fresh charge of catalyst on stream should be as l-ow as possible (commensurate with the maintenance of adequate perpass conversion levels) since the lower the starting temperature the longer will be the duration `of the said onstream period. For any given conversion, the permissible starting temperature is a function of catalyst activity inasmuch as the more active catalysts (i.e., those capable of effecting a relatively high per-pass conversion under given operating condi-tions) permit the unit to be placed on stream at lower startin g temperatures than would otherwise be the case. Preferred initiating temperatures are inthe range bf from about 350 to 650 F.

The etiluent from the hydrocracking reactor may be Worked up in any convenient fashion. Thus, a gas re-` cycle stream rich in hydrogen iscustomarily separated in a high pressure gas-liquid separation zone, which stream can be recycled for admixture with the feed passing over the hydrocracking catalyst. Thereafter, C1-C3 or C1-C4 products are separated in a gas-liquid separation zone operated at lower pressures, thus leaving a normally liquid eluent portion from which fractions boiling in the gasoline range can be recovered. Said fractions may be used as gasoline components, or more preferably, their octane ratings can be greatly improved by passing the heavier portions thereof (e.g., those boiling within a range of from about 165 to 400 F.) through a catalytic reformer under reforming conditions. Products in the hydrocracker effluent boiling above the desired end point of the gasoline fractions can be employed in whole or in part for jet or other fuel purposes, or they may be recycled backover Vthe hydrocracking catalyst. Thus, in one embodiment of the invention, fractions boiling below about 350-400 F. are recovered for use as gasoline blending stocks (after reforming, if desired), a next higher boiling fraction having an end point of about 550 to 600 is diverted for jet fuel purposes, and any remaining bottoms are recycled over the hydrocracking catalyst.

' Reference has been made above to the fact that appropriate, higher boiling gasoline fractions recovered `from the hydrocracker effluent can be upgraded in octane number by a reforming treatment. In carrying out the lat-ter, the feed is passed, along with added hydrogen, over a catalyst in a reforming zone under conventional reforming conditions, representative reforming temperatures and .pressures being from about 850-l000 F. and from about 200 to 900 p.s.i.g. The catalyst employed in the reforming zone can be any one or more of the various materials which are now available for effecting reforming opera- 101 tions. Oe such catalyst comprises from about 8 to 12% molybdenum oxide disposed on an alumina support. Another contains from about 0.1 to 1% by weight of catalytic platinum dispersed on an alumina support along with small amounts of halogens.

The manner lin which the present invention is practiced can be illustrated by reference to the figure of the appended drawing which is a simplified fl-ow scheme of a refinery unit suitable for use in. the process.

In the drawing, a reduced petroleum crude or other residual oil is shown as being passed through line 10 to a coker unit 11 from which coke is discharged through line 12, while gaseous hydrocarbon products formed during c'oking are passed via line 13 to a stabilizer column 14. Normally gaseous products present in the coker effluent stream are taken overhead from the stab-ilizer through line 15, theremaining, normally liquid products being passed via line 416 to fracti-ona'ting column 17 from which light and heavy gasoline streams are discharged Athrough lines 18 and 19. A light coker gas oil stream is recovered from the fractiona-tor and passed through line 20 to a hydroner 21, while heavy coker gas oil is taken as a bottoms stream through line 22.` Said heavy gas oil stream may be recycled to the colier through line 23, if desired; alternatively, it may be passed through line V24 to the hydroiiner 21 or be used as feed to a cracking unit 25.

In the hydroner 2.1 which, Iin addition to light gas oil and any heavy gas oil supplied thereto, may also receive other streams such as a straight-run gas oil (line 26) or a light cycle oil (line 27) produced in cracker 25, the hydrocarbon feed is admixed with hydrogen introduced via line 28 and passed over a hydrogenation catalyst under conditions of elevated temperature and pressure effective to selectively decompose nitrogen-containing and sulfur-containing compounds which may be present. The product from the hydroner is discharged through line 28 and, following admixture with water introduced through line 29, is discharged into a gas-liquid separator 30from which a hydrogen-rich recycle stream is taken overhead through line 31 for recycle to the hydrofiner along with make-up hydrogen from line 32. In the separator 30 there is formed an upper hydrocarbon layer which is'drawn off through line 33 and passed to stripper 34, and a lower, aqueous layer which contains various of the ni-trogen-` and sulfur-containing decomposition products formed in the hydroiner, said lower -layer being discharged as waste.

In the stripper 34, which is operated at lower pressure than the separator 30, a gaseous stream separates and is taken overhead via line 35, while the remaining, denitried hydrocarbon stream is passed through line 36 and, after being admixed with incoming hydrogen from line 37, is passed over the catalyst in the hydrocracker 38 under conditions of elevated temperature and pressure.

The effluent stream fro-m the hydrocracker 38 is passed via line 39 to a high pressure gas-liquid separator 40 from which a hydrogen-rich stream is taken overhead through line 41 for return to the hydrocracker along with makeup hydrogen from line 32, with the remaining efuent portions being passed through line 42 to a low pressure gas-liquid separator 43 from which remaining hydrogen, together with C1 and C2 products, are taken overhead as gas through line 44. The remaindery of the stream is directed through line 45 into a fractionating column 46 which serves to remove C3-and C4 products (line 47), leaving a C5-I- effluent stream which is passed through line 48 to a fractionating column 49 from which a light gasoline'product stream is taken overhead through line 50, While a heavy gasoline cut (for example, one boiling from about to 400 F.) is passed through line 51 to a reformer 52. Effluent portions from the hydrocracker 38 boiling above the heavy gasoline cut are taken as bottoms vthrough line 53 and may either be recycled to the hydrocracker through line 54 or be diverted (line 55) to jet fuel or other renery end uses.

In the reformer 52, the heavy gasoline feed is admixed with hydrogen from line 56 and passed over a platinumon-alumina or other reforming catalyst at generally higher temperatures and lower pressures than those prevailing over the hydrocracking catalyst` whereby appropriate feed compounds are dehydrogenated or cyclicized and dehydrogenated to form the corresponding aromatic compounds. The efuent stream from the reformer, which, as a result of the reactions taking place therein, contains appreciable quantities of hydrogen over and above that introduced via line 56, is discharged through line 57 to a gas liquid separator 58 from which a hydrogen rich stream is taken overhead through line 32, while the balance of the stream passes through line 59 to a stabilizer 60. Normally gaseous products are removed from the stabilizer through line 61, while a gasoline fraction, now considerably upgraded in octane rating, is discharged through line 62.

Returning to the cracker 25 which, as noted above, may be supplied (among other feed sources) with heavy coker gas oil feed, the conditions are such that the feed stream is cracked to lighter boiling product fractions either under the inuence of heat alone or of heat and an appropriate cracking catalyst. The eftluent stream from the cracker is discharged through line 63 to fractionator 64 from which a gas stream may be taken off overhead through line 65, with light and heavy gasoline fractions .being recovered through line 66 and 67, respectively, said heavy gasoline being passed, in some instances, to the reformer 52. A light cycle oil cut normally boiling within arange from about 350 to 600 F., or somewhat higher, is recovered through line 27 and passed to the hydroner 21 for la-ppropriate denitrication treatment. A heavy cycle oil stream is removed as bottoms from fractionator 64 and may be recycled to the cracker 25 through line 68 or be diverted through line 69 to other appropriate refinery uses.

The advantages to be gained by a practice of the present invention are illustrated by the data of the following example.

EXAMPLE I In this operation a crude petroleum residual stock representing a mixture of fractions from various California crude oils was passed through a coking unit of the delayed type. The feed supplied to the coker boiled above about 950 F. and had an API gravity of about 6.1 and a viscosity of approximately 20,000 SSU at 210 F., its Conradson carbon and sulfur contents being of the order of 14.9 and 2.5 wt. percent, respectively. As a result of the coking operation, there was recovered a light coker gas oil having the following inspections.

Gravity, API 24.8

This light gas oil was denitried by first passing the same, along with 6000 s.c.f. H2 per barrel, at 720 p.s.i.g., 730 F. and 0.5 LHSV, over a hydroning catalyst comprising molybdenum oxide (18.6 wt. percent Mo) with nickel sulfide (5.5 wt. percent Ni) and cobalt sulde (2.1 wt. percent Co) on an alumina support. The product recovered from this denitrication step was then processed a second time over the catalyst under the same conditions, except that the pressure was raised to 1200 p.s.i.g. Total 'hydrogen consumption was approximately 1000 s.c.f/bbl. of feed, and the hydrotlned product, following removal of NH3, and HZS, and other gases, had the following inspections.

Gravity, API 35.1

The foregoing hydroned stock was then hydrocracked by passing the same, along with 6900 s.c.f. Hz/ bbl. feed, at 1200 p.s.i.g., 626 F. and 0.8 LHSV over a catalyst cornprising nickel sulfide (6 wt. percent nickel) on a synthetic silica-alumina cracking support containing about silica, said catalyst being Catalyst A as described above. Under these conditions, there was obtained a per-pass conversion of 58.8 percent to synthetic products boiling below 400 F. with all elluent boiling above 400 F. being recycled over the catalyst. The operation was consumptive of 1063 s.c.f Hg/bbl. feed converted to said products. The weight and volume percent yields of the various synthetic products, based on the feed converted thereto, was as follows:

Wt. Vol percent percent The C5-180 F. cut referred to above was estimated to have an octane number of approximately 100, F-1-l-3 ml. TEL. The -400 F. cut had the following inspections.

The above-described 180-400 F. cut had an F-l clear octane rating of about 69. However, it is found that the octane rating thereof could be raised to a value of approximately 103 (F-l-I-3 ml. TEL) by passing the cut, along with 6000 s.c.f. Hz/bbl. of feed, at a temperature of 920 F. and 500 p.s.i.g. over a reforming catalyst comprising 0.75% platinum disposed on an alumina support at an LHSV of 2.

A material balance based on a run conducted under 13 conditions generally similar to those set forth in this example has been given above in Table III.

We claim:

1. A process for converting hydrocarbon feeds of a heavy, residual character to gasoline and other product fractions boiling below said feeds, said process comprising passing the feed through a coking unit under coking conditions; recovering from the coking unit a light, ,nitro gen-containing oil and a heavy gas oil; passing said heavy gas oil through a cracking unit under cracking conditions; separating the effluent from the cracking unit into fractions including a light, nitrogen-containing cycle oil boiling within a range of from about 350 to 650 F. and a heavier fraction; subjecting the light gas oil from the coking unit and said light cycle oil fraction from the cracking unit to denitrication treatment wherein said oils are passed, along with added hydrogen, over a hydrogenation catalyst under elevated temperature and pressure conditions effective to crack nitrogen compounds contained therein, and recovering denitried gas oil and cycle oil fractions each containing a total of less than 10 p.p.m. of nitrogen; and contacting said denitried fractions in a hydrocracking zone, along with at least 2000 s.c.f. H2/bbl. thereof, with an acidic catalyst having hydrogenating characteristics and high cracking aetivity at temperatures of from about 350 to 700 F. and at pressures of at least 400 p.s.i.g., whereby a substantial 14 portion of said fractions is converted to gasoline product fractions.

2. The process of claim 1 wherein, of the last mentioned gasoline product fractions, those of higher boiling point are passed, along with added hydrogen, over a reforming catalyst under reforming conditions.

3. The process of claim 1, wherein said denitried fractions are contacted with said catalyst in said hydrocracking zone at hydrocracking zone initiating temperatures of from about 350 to 650 F. when said hydrocracking zone is placed on stream, and wherein said temperature is gradually raised during the on-stream period as necessary to maintain the desired per pass conversion.

4. The process of claim 1, wherein said denitried fractions each contain less than 2 p.p.m. of nitrogen.

5. The process of claim 1, wherein said heavier frac# tion from said cracking unit is recycled to extinction in that unit.

References Cited in the file-of this patent UNITED STATES PATENTS 2,727,853 Hennig Dec. 20, 1955 2,839,450 Oettinger June 17, 1958 2,937,134 Bowles May 17, 1960 3,008,895 Hansford et al. Nov. 14, 1961 UNITED STATES PATENT OFFICE CERTIFICATE OF CORRECTION Patent No. 3,072,560 January-8,- 1963 Norman J. Paterson et al..

It is hereby certified that error appears in the above numbered patent requiring correction and that the said Letters Patent should read as corrected below Column 7, lines 6 and 7, after. 'ISerial No. 794,109, ziled February 18, 1959" strike ou-.t'T and now `abandoned" Signed and sealed this 12thl day of May-M1964.

TI'EAL) yfntest:

ERNEST W. SWIDER EDWARD J. BRENNER E1 ttesting Ufficer Commissioner of Patents 

1. A PROCESS FOR CONVERTING HYDROCARBON FEEDS OF A HEAVY, RESIDUAL CHARACTER TO GASOLINE AND OTHER PRODUCT FRACTIONS BOILING BELOW SAID FEEDS, SAID PROCESS COMPRISING PASSING THE FEED THROUGH A COKING UNIT UNDER COKING CONDITIONS; RECOVERING FROM THE COKING UNIT A LIGHT, NITROGEN-CONTAINING OIL AND A HEAVY OIL; PASSING SAID HEAVY GAS OIL THROUGH A CRACKING UNIT UNDER CRACKING CONDITIONS; SEPARATING THE EFFLUENT FROM THE CRACKING UNIT INTO FRACTIONS INCLUDING A LIGHT, NITROGEN-CONTAINING CYCLE OIL BOILING WITHIN A RANGE OF FROM ABOUT 350 TO 650*F. AND A HEAVIER FRACTION; SUBJECTING THE LIGHT GAS OIL FROM THE COKING UNIT AND SAID LIGHT CYCLE OIL FRACTION FROM THE CRACKING UNIT TO DENITRIFICATION TREATMENT WHEREIN SAID OILS ARE PASSED, ALONG WITH ADDED HYDROGEN, OVER A HYDROGENATION CATALYST UNDER ELEVATED TEMPERATURE AND PRESSURE CONDITIONS EFFECTIVE TO CRACK NITROGEN COMPOUNDS CONTAINED THEREIN, AND RECOVERING DENITRIFIED GAS OIL AND CYCLE OIL FRACTIONS EACH CONTAINING A TOTAL OF LESS THAN 10 P.P.M. OF NITROGEN; AND CONTACTING SAID DENITRIFIED FRACTIONS IN A HYDROCRACKING ZONE, ALONG WITH AT LEAST 2000 S.C.F. H2/BBL. THEREOF, WITH AN ACIDIC CATALYST HAVING HYDROGENATING CHARACTERISTICS AND HIGH CRACKING ACTIVITY AT TEMPERATURES OF FROM ABOUT 350 TO 700*F. AND AT PRESSURES OF AT LEAST 400 P.S.I.G., WHEREBY A SUBSTANTIAL PORTION OF SAID FRACTIONS IS CONVERTED TO GASOLINE PRODUCT FRACTIONS. 